|Publication number||US5188725 A|
|Application number||US 07/669,720|
|Publication date||Feb 23, 1993|
|Filing date||Mar 15, 1991|
|Priority date||Mar 15, 1991|
|Publication number||07669720, 669720, US 5188725 A, US 5188725A, US-A-5188725, US5188725 A, US5188725A|
|Inventors||Mohsen N. Harandi|
|Original Assignee||Mobil Oil Corporation|
|Export Citation||BiBTeX, EndNote, RefMan|
|Patent Citations (12), Referenced by (43), Classifications (13), Legal Events (4)|
|External Links: USPTO, USPTO Assignment, Espacenet|
The present invention relates to a multi-stage process for cracking hydrocarbons and preparing ethers. More particularly, it relates to a technique for utilizing zeolite catalyst used in preparing ethers as makeup or equilibrium catalyst for large scale fluidized catalytic cracking (FCC) processes.
This invention relates to a catalytic technique for cracking heavy petroleum stocks and converting olefin gas to valuable ethers. In particular, it provides a continuous integrated process for etherifying olefinic light gas byproduct of FCC cracking to produce C5 + ethers. Isobutylene and/or isoamylene containing streams, byproducts of petroleum cracking in a fluidized catalytic cracking (FCC) unit, may be upgraded to ethers by contact with a solid acid catalyst, such as crystalline medium pore siliceous zeolite catalyst.
Recently it has been found that ethers can be produced from FCC iso-olefins by catalytic conversion in a fluidized bed of solid medium pore acid zeolite catalyst. Such a fluidized bed operation typically requires oxidative regeneration of coked catalyst to restore zeolite acidity for further use, while withdrawing spent catalyst and adding fresh acid zeolite to maintain the desired average catalyst activity in the bed. This technique is particularly useful for upgrading FCC light olefinic gas, which usually contains significant amounts of C4 olefins, including isobutene.
Economic benefits and increased product quality can be achieved by integrating the FCC and etherification units in a novel manner. It is the primary object of this invention to eliminate the etherification catalyst regeneration system which results in significant process investment saving and improved process safety. Another object of this invention is to eliminate the etherification catalyst regeneration which results in significant process investment/operating cost saving. Another object of the present invention is to further extend the usefulness of the medium pore acid zeolite catalyst used in the ether reaction by withdrawing a portion of partially deactivated and coked zeolite catalyst and admixing the withdrawn portion with cracking catalyst in a primary FCC reactor stage. The catalyst withdrawn from the etherification unit operations can be sent directly to the FCC reactor or regenerator; however, it is also feasible to employ the catalyst in other intermediate unit operations, such as olefin upgrading. Prior efforts to increase the octane rating of FCC gasoline by addition of zeolites having a ZSM-5 structure to large pore cracking catalysts have resulted in a small decrease in gasoline yield, increase in gasoline quality, and increase in light olefin byproduct.
Recent efforts have been made in the field of gasoline blending to increase gasoline octane performance without the addition of deleterious components such as tetraethyl lead and benzene. It has been found that lower molecular weight unsymmetrical ethers, such as methyl t-butyl ether (MTBE) and t-amyl methyl ether (TAME) can be added to C5 -C10 hydrocarbon-containing gasoline products. Conventional etherification processing uses as catalyst a macroreticular cation exchange resin in the hydrogen form. An example of such a catalyst is "Amberlyst 15". A resin catalyst gives a high conversion rate but is unstable at elevated temperatures (above about 90° C.). When overheated, the resin catalyst releases sulfonic and sulfuric acids. In addition leaching of acid substances from the resin catalyst even at normal operating temperatures causes a reverse reaction--decomposition of ether products to starting materials--to occur upon distillation of ether product. Overall yield is thereby significantly decreased (see U.S. Pat. No. 4,182,913 to Takesono et al).
Etherification reactions conducted over a resin catalyst such as "Amberlyst 15" are usually conducted in the liquid phase below a temperature of about 90° C. and at a pressure of about 150-200 psig. Equilibrium is more favorable at lower temperatures but the reaction rate decreases significantly. Also excess methanol appears to be required to achieve acceptable selectivity. Some recent efforts in the field of etherification reactions have focused on the use of acid medium-pore zeolite catalyst for highly selective conversion of iso-olefin and alcohol starting materials. Examples of such zeolite catalysts are ZSM-4, ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-50, MCM-22 and zeolite Beta, as disclosed in copending U.S. patent application Ser. No. 07/495,667 (Harandi et al/U.S. Pat. No. 5,015,782). Due to lower acidity as compared to resin catalysts, the zeolites need to be employed at higher reaction temperature to achieve conversion rates substantially equivalent to resin catalysts. These solid acid catalyst particles are much more thermally stable than resin catalyst, are less sensitive to methanol-to-isobutene ratio, give no acid effluent, and are easily and quickly regenerated (see Chu et al, "Preparation of Methyl tert-Butyl Ether (MTBE) over Zeolite Catalysts", Industrial Engineering and Chemical Research).
Developments in zeolite catalysis and hydrocarbon conversion processes have created interest in utilizing olefinic feedstocks for producing C4 -C5 tertiary olefins, gasoline, etc. In addition to basic chemical reactions promoted by zeolite catalysts having a ZSM-5 structure, a number of discoveries have contributed to the development of new industrial processes for improved FCC operation to enhance iso-olefin production.
It is an object of the present invention to provide a process and apparatus for continuous operation in preparation of alkyl ethers from lower alkanol and C4 olefin with a conventional acid resin catalyst whereby the resin catalyst is protected from impurities such as nitrogen compounds, metals, and coke. It is a further object to use catalyst from such etherification unit operations as makeup for FCC units.
An improved process has been found for producing ether by catalytic contact of etherification feedstock comprising alkene, alkanol, ether precursors or mixtures thereof in the presence of at least one deactivating impurity, with thermally stable solid material having acid catalytic activity under etherification conditions.
The preferred improvement comprises a multi-stage process for increasing production of high octane gasoline range hydrocarbons from cracked FCC petroleum feedstock, including the steps of: contacting the feedstock in a primary fluidized catalyst reaction stage with a mixed catalyst system which comprises finely divided particles of a first large pore cracking catalyst component and finely divided particles of a second medium pore siliceous zeolite catalyst component under cracking conditions to obtain a product comprising intermediate gasoline and distillate range hydrocarbons, and an olefinic gas rich in C4 + iso-olefin; separating the olefinic gas and contacting at least a fraction of said olefins together with lower aliphatic alcohol with particulate catalyst solids consisting essentially of acid medium pore siliceous zeolite catalyst in a secondary fluidized bed reaction stage under etherification reaction conditions, thereby depositing about carbonaceous material onto the particulate zeolite catalyst to obtain a coked equilibrium catalyst; withdrawing a portion of partially deactivated equilibrium particulate zeolite catalyst from the secondary reaction stage; and adding a said withdrawn coked equilibrium zeolite catalyst to the primary fluidized reaction stage for conversion of crackable petroleum feedstock. Employing this technique, catalyst makeup of a primary stage fluidized catalytic cracking unit and a secondary stage etherification reactor is balanced. For instance, catalyst flow rates per day are adjusted so that about 1 to 10 percent by weight of fresh cracking catalyst based on total amount of catalyst present in the primary fluidized bed reaction stage is added to the primary reaction stage; about 0.5 to 100 percent by weight fresh zeolite catalyst based on total amount of catalyst present in the secondary fluidized bed reaction stage is added to the secondary reaction stage; and about 0.5-100 percent by weight of partially deactivated zeolite catalyst based on total amount of catalyst present in the secondary reaction stage is withdrawn from the secondary reaction stage and added to the primary fluidized bed reaction stage.
FIG. 1 is an overall flow sheet depicting FCC and ether units and their processing relationships: and
FIG. 2 is a schematic diagram of a preferred embodiment of the etherification unit of the present process, showing major operating units and flow of reactants and chemical products.
FIG. 1 illustrates a process scheme for practicing the present invention. The flow of chemicals beginning with the heavy hydrocarbons feed at line 1 is schematically represented by solid lines. The flow of catalyst particles is represented by dotted lines. Chemical feedstock passes through conduit 1 and enters the first stage fluidized bed cracking reactor 10. The feed can be charged to the reactor with a diluent such as hydrocarbon or steam. Deactivated catalyst particles are withdrawn from fluidized bed reaction zone 10 via line 3 and passed to catalyst regeneration zone 40, where the particles having carbonaceous deposits thereon are oxidatively regenerated by known methods. The regenerated catalyst particles are then recycled via line 5 to reaction zone 10. It may be feasible to utilize regenerated catalyst from the FCC unit operations in ether unit processing, especially wherein a substantial amount of large pore (e.g.- zeolite Y) catalyst can be beneficial in a pre-etherification guard chamber unit.
A portion of secondary stage catalyst is sent via conduit 37 to first fluid bed reaction zone 10. Fresh medium pore zeolite catalyst can be admixed with the regenerated catalyst as by conduit 39. Also, fresh medium pore zeolite catalyst is added to etherification reaction zone 30 via conduit 20.
Cracked product from the FCC reaction zone 10 is withdrawn through conduit 2 and passed to a main fractionation tower 4 where the product is typically separated into a light gas stream, a middle stream, and a bottoms stream. The middle stream is recovered via conduit 12 and the bottoms stream is withdrawn through conduit 11. The light gas stream is withdrawn through conduit 6 and enters gas plant 8 for further separation. A middle fraction is drawn from the gas plant via conduit 14 and a heavy fraction is withdrawn via conduit 13. A stream comprising C4+ iso-olefin is withdrawn via conduit 7, with optional water washing and enters etherification unit 30 where the stream contacts zeolite solid catalyst particles in a turbulent regime fluidized bed or the like to form the desired ether product. Lower alkanol enters the reactor of unit 30 concurrently with olefin. Ether rich (e.g.- MTBE) product is removed from the etherification unit 30 through conduit 9.
The catalyst inventory in the FCC reactor preferably comprises predominantly zeolite Y which is impregnated with one or more rare earth elements (REY). This large pore cracking catalyst is preferably combined in the FCC reactor with the ZSM-5 withdrawn from the etherification reactor to obtain a mixed FCC cracking catalyst which provides a gasoline yield having improved octane number and an increased yield of lower molecular weight iso-olefins.
Advantageously, the catalyst flow rates per day are adjusted so that about 1 to 10 percent by weight of fresh cracking catalyst based on total amount of catalyst present in the primary fluidized bed reaction stage is added to the primary reaction stage; about 0.5 to 100 percent by weight fresh zeolite catalyst based on total amount of catalyst present in the secondary fluidized bed reaction stage is added to the secondary reaction stage; and about 0.5-100 percent by weight of partially deactivated zeolite catalyst based on total amount of catalyst present in the secondary reaction stage is withdrawn from the secondary reaction stage and added to the primary fluidized bed reaction stage to increase octane by 0.2-5 Research (base 92 Research).
Catalyst inventory in the fluidized catalytic cracking unit may be controlled so that the ratio of cracking catalyst to the added zeolite etherification catalyst is about 1:1 to about 50:1. The fresh catalyst for the etherification unit and the FCC unit has adequate acidity to effect etherification.
In a preferred example, the total amount of fluidized catalyst in the FCC reactor is about ten times as much as the amount of fluidized catalyst in the etherification reactor. To maintain equilibrium catalyst activity in the FCC reactor, fresh Y zeolite catalyst particles are added in an amount of about 1 to 2 percent by weight based on total amount of catalyst present in the FCC reactor. Spent cracking catalyst is then withdrawn for subsequent disposal from the FCC regenerator in an amount substantially equivalent to the combination of fresh REY zeolite catalyst and partially deactivated ZSM-5 catalyst which is added to the reactor.
In a typical example of the present process, an FCC reactor is operated in conjunction with an etherification reactor. The catalyst flow rates per day are adjusted so that about 1.25 percent by weight of fresh large pore zeolite cracking catalyst based on total amount of catalyst present in the FCC reactor is added to the FCC reactor; about 10.0 percent by weight fresh zeolite ZSM-5 catalyst based on total amount of catalyst present in the ether unit zeolite reactor is added to the ether unit; and about 10.0 percent by weight of zeolite ZSM-5 catalyst based on total amount of catalyst present in the ether reactor is withdrawn and added to the catalyst inventory of the FCC reactor. The gasoline range hydrocarbons obtained from the FCC reactor have an increased octane rating (using the R+M/2 method, where R=research octane number and M=motor octane number) of 0.7.
In conventional fluidized catalytic cracking processes, a relatively heavy hydrocarbon feedstock, e.g., a gas oil, is admixed with hot cracking catalyst, e.g., a large pore crystalline zeolite such as zeolite Y, to form fluidized suspension. A fast transport bed reaction zone produces cracking in an elongated riser reactor at elevated temperature to provide a mixture of lighter hydrocarbon crackate products. The gasiform reaction products and spent catalyst are discharged from the riser into a solids separator, e.g., a cyclone unit, located within the upper section of an enclosed catalyst stripping vessel, or stripper, with the reaction products being conveyed to a product recovery zone and the spent catalyst entering a dense bed catalyst regeneration zone within the lower section of the stripper. In order to remove entrained hydrocarbon product from the spent catalyst prior to conveying the latter to a catalyst regenerator unit, an inert stripping gas, e.g., steam, is passed through the catalyst where it strips such hydrocarbons conveying them to the product recovery zone. The fluidized cracking catalyst is continuously circulated between the riser and the regenerator and serves to transfer heat from the latter to the former thereby supplying the thermal needs of the cracking reaction which is endothermic. Particular examples of such catalytic cracking processes are disclosed in U.S. Pat. No. 4,927,526 (Anderson et al) and U.S. patent application Ser. No. 07/495,667, filed Mar. 19, 1990 (Harandi/Docket 5712), incorporated herein by reference.
Several of these processes employ a mixture of catalysts having different catalytic properties as, for example, the catalytic cracking process described in U.S. Pat. No. 3,894,934 which utilizes a mixture of a large pore crystalline zeolite cracking catalyst such as zeolite Y and shape selective medium pore crystalline metallosilicate zeolite such as ZSM-5. Each catalyst contributes to the function of the other to produce a gasoline product of relatively high octane rating.
A fluidized catalytic cracking process in which a cracking catalyst such as zeolite Y is employed in combination with a shape selective medium pore crystalline include the synthetic faujasite zeolites X and Y with particular preference being accorded zeolites Y, REY, USY and RE-USY.
The shape selective medium pore crystalline zeolite catalyst can be present in the mixed catalyst system over widely varying levels. For example, the zeolite of the second catalyst component can be present at a level as low as about 0.01 to about 1.0 weight percent of the total catalyst inventory (as in the case of the catalytic cracking process of U.S. Pat. No. 4,368,114). In the present invention it can represent as much as 50 weight percent of the total catalyst system.
The catalytic cracking unit is preferably operated under fluidized flow conditions at a temperature within the range of from about 480° C. to about 735° C., a first catalyst component to charge stock ratio of from about 2:1 to about 15:1 and a first catalyst component contact time of from about 0.5 to about 30 seconds. Suitable charge stocks for cracking comprise the hydrocarbons generally and, in particular, petroleum fractions having an initial boiling point range of at least 205° C., a 50% point range of at least 260° C. and an end point range of at least 315° C. Such hydrocarbon fractions include gas oils, thermal oils, residual oils, cycle stocks, whole top crudes, tar sand oils, shale oils, synthetic fuels, heavy hydrocarbon fractions derived from the destructive hydrogenation of coal, tar, pitches, asphalts, hydrotreated feedstocks derived from any of the foregoing, and the like. As will be recognized, the distillation of higher boiling petroleum fractions above about 400° C. must be carried out under vacuum in order to avoid thermal cracking.
Since mixed olefinic ether feedstock may contain many impurities, even after treatment in a "Merox" unit and a water wash, the solid acid zeolite etherification catalyst can become highly contaminated after a period of on-line contact with feedstock. Some of the impurities which are absorbed on the zeolite particles are: nitrogen compounds; metals such as Al, Fe, Na and Mg; and oligomers of olefins and diolefins, such as butadiene, isoprene and cyclopentadiene. Diolefinic compounds and other related hydrocarbons are deposited as coke on the surface and interstices of the zeolite and/or resin catalytic particles. It is therefore an objective of the present process to remove feedstock impurities in a guard chamber or first reaction zone prior to or concurrently with the preparation of ethers. The first etherificaiton reaction zone preferably comprises at least one reactor containing fluidizable catalyst. The effluent of this reactor may contain catalyst solids, which can be recovered by filtration prior to return to the reactor and/or FCC unit.
In the preferred embodiment, the first reaction etherification zone contains one or more reactors containing a fluidized bed of finely divided zeolite catalyst. The preferred solid acid catalyst particles are aluminosilicate zeolites selected from ZSM-5, ZSM-11, ZSM-35, ZSM-50, MCM-22 and zeolite Beta; however, it is feasible to employ at least a minor portion of large pore cracking catalyst e.g.- REY). It is within the scope of the present process and apparatus to adjust the number and types of reactors which contain acid zeolite catalyst in order to optimize both product yield and overall energy consumption as would be practiced by one skilled in the art. If a zeolite guard chamber is employed for etherification feed pretreatment, a temperature up to about 50° C. (75°-120° F.) is preferred for contact with zeolite, especially wherein no alcohol is cofed to the guard chamber, of particular importance in treating C5 + olefins. and The mixed olefinic feedstock can contain a significant amount of impurities; however, if desired, the step of washing the feedstream with water can be eliminated.
The optional second etherification reaction zone contains an acid resin catalyst which is preferably a macroreticular polystyrene sulfonic acid resin catalyst. In a preferred embodiment the second reaction zone contains a catalytic distillation column containing polystyrenesulfonic acid resin catalyst in a plurality of fixed bed catalysis-distillation units located in the upper half of the distillation column. The reaction section containing resin catalyst is preferably operated at a temperature about 10° to 30° C. lower than the temperature of the first reaction zone.
In an alternate embodiment, the second reaction zone is not a catalytic distillation column, but rather a single reactor or plurality of reactors. Reactor configuration can take many forms, for example, fixed bed, stirred slurry (see U.S. Pat. No. 3,940,450 to Lee, incorporated herein by reference), swing or ebullated bed. It is within the scope of the present process to employ for the second reaction zone any reactor configuration for sequencing acceptable to the skilled engineer. The present invention contemplates that an acid resin catalyst be employed following a regenerable etherification catalyst, preferrably in the second reaction zone. In a preferred embodiment, the resin catalyst is "Amberlyst 15".
A signicant improvement is found by adding a preliminary step of contacting the olefin and alcohol reactants in the liquid phase with oxidatively regenerable solid acid catalyst particles in a preliminary reaction zone under partial etherification conditions to produce an intermediate stream comprising tert-alkyl ether and unreacted olefin and alcohol, said intermediate stream being substantially free from impurities which reduce catalyst activity. In a preferred embodiment the olefinic feedstock comprises isobutene in an amount of at least about 10 wt. %. Preferably the acid catalyst is aluminosilicate having the structure of ZSM-5, zeolite Beta and/or zeolite Y and is contained in a fluid bed reactor for ease of removal from contact with reactants. Once removed from on-line activity, the acid catalyst is passed to the FCC unit for conversion of heavy hydrocarbons to lighter product, including etherifiable lower olefins.
Although the preferred alcohol is methanol, suitable substitutes include ethanol or isopropanol (isopropyl alcohol). Of course, use of these substitutes will yield different ether products. It is within the scope of the present process to employ a mixture of lower molecular weight alcohols. Although C4 hydrocarbons containing isobutene, is the preferred hydrocarbon feed, other C5-C7 tertiary olefins such as 3-methyl-2-butene can be etherified in the present process.
Referring to FIG. 2, a pre-washed C4 + aliphatic hydrocarbon feedstock stream 110 containing isobutane is fed with methanol stream 112 to a Stage I etherification reactor 120 for contact with a solid zeolite from supply conduit 122 and the mixed alcoholic C4 + hydrocarbon feedstream contacts the solid catalyst within this reaction zone at predetermined reaction zone conditions of temperature and pressure to convert at least a portion of the feedstream to MTBE. It is understood that fluidization can be effective in liquid-solid slurry or a turbulent gas-solid catalyts bed. Impurities present within the feedstock are effectively removed from the partially converted feedstream by the solid acid catalyst. Catalyst particles entrained in the reactor effluent can be recovered by filtration. Catalyst withdrawn from Stage I reactor via conduit means 123 is transferred to the FCC unit operations via conduit 37, as discussed above.
Intermediate stream 124 containing MTBE and unreacted C4 hydrocarbons and alcohol is withdrawn from reaction zone 120 and enters Stage II catalytic distillation column 130. The temperature of the intermediate stream may be adjusted prior to entering the distillation column. In distillation column 130 a substantial portion of unreacted C4 hydrocarbons and alcohols are converted to MTBE over a polystyrenesulfonic acid resin catalyst such as "Amberlyst 15". Etherification over resin catalyst is carried out preferably at a temperature of about 37° to 75° C. and a pressure of about 10 to 350 psig. In a preferred embodiment acid resin catalyst is placed in an upper rectifying section 32 of a debutanizer column used for stabilizing the ethers. A product stream comprising MTBE can be withdrawn from a lower portion of distillation column 130 by line 134. Unreacted light gases are removed as by line 136. The etherification unit operations are described in U.S. patent application Ser. No. 07/339,466, filed Apr. 17, 1989 (Harandi/U.S. Pat. No. 5,000,837), incorporated herein by reference. At least a portion of alcohol can be added to the secondary reactor 130 via line 124.
To illustrate the common problem of catalyst poisoning when a polysulfonic acid resin catalyst is employed in the etherification process, MTBE resin catalyst unit is operated in a continuous fashion for a period of six months. Isobutene containing hydrocarbon feed is purified in a "Merox" unit and water-washed prior to entering the MTBE reactor. Conversion decreases from 93% to 52% during the six month period. Analysis identifies the contaminants on the resin catalyst. The major contaminants are nitrogen compounds, which are responsible for about 60% of the catalyst deactivation. The concentration of nitrogen on the deactivated resin catalyst is about 25×103 ppm. Metals such as Al, Fe, Na and Mg account for about 10% of the deactivation. The source of such metals is mainly from the water wash tower. The concentration of the metals on the deactivated catalyst is about 15×102 ppm. The third type of contaminant is coke. Coke is formed on the resin catalyst due to the presence of such compounds as cyclopentadiene and isoprene in the hydrocarbon feedstock. Continuous monitoring of the feedstock is necessary to control particularly the diolefinic C5 hydrocarbon content. One of the advantages of the present process is that coke formation occurs primarily on the zeolite catalyst.
It is also observed that acetone and nitrile compounds are major contaminants in the hydrocarbon feedstocks which have been water washed. For example, a feed sample may contain 190 ppm acetone, 3 ppm acetonitrile and 16 ppm propionitrile. An advantage of the present process is that the hydrocarbon feedstock does not have to be water washed.
While the invention has been described by specific examples and embodiments, there is no intent to limit the inventive concept except as set forth in the following claims.
|Cited Patent||Filing date||Publication date||Applicant||Title|
|US4826507 *||Dec 8, 1987||May 2, 1989||Mobil Oil Corporation||Integrated etherification and oxygenates to gasoline process|
|US4827046 *||Apr 11, 1988||May 2, 1989||Mobil Oil Corporation||Extraction of crude methanol and conversion of raffinate|
|US4830635 *||Dec 8, 1987||May 16, 1989||Mobil Oil Corporation||Production of liquid hydrocarbon and ether mixtures|
|US4835329 *||Dec 8, 1987||May 30, 1989||Mobil Oil Corporation||Process for producing high octane gasoline|
|US4854939 *||Dec 8, 1987||Aug 8, 1989||Mobil Oil Corporation||Aromatization and etherification process integration|
|US4925455 *||Mar 17, 1989||May 15, 1990||Mobil Oil Corporation||Process for the etherification of linear and branched olefins|
|US4990712 *||May 18, 1990||Feb 5, 1991||Mobil Oil Corporation||Integrated cracking, etherification and olefin upgrading process|
|US5001292 *||Oct 16, 1989||Mar 19, 1991||Mobil Oil Corporation||Ether and hydrocarbon production|
|US5011506 *||Sep 29, 1989||Apr 30, 1991||Mobil Oil Corporation||Integrated etherification and alkene hydration process|
|US5013329 *||Mar 31, 1989||May 7, 1991||Mobil Oil Corporation||Conversion of light hydrocarbons to ether rich gasoline|
|US5047070 *||May 30, 1989||Sep 10, 1991||Mobil Oil Corporation||Integrated process for production of gasoline and ether from alcohol with feedstock extraction|
|US5120881 *||Mar 30, 1990||Jun 9, 1992||Exxon Chemical Patents Inc||Removal of nitrogenous components of a hydrocarbon feedstream|
|Citing Patent||Filing date||Publication date||Applicant||Title|
|US5271835 *||May 15, 1992||Dec 21, 1993||Uop||Process for removal of trace polar contaminants from light olefin streams|
|US5569790 *||May 3, 1994||Oct 29, 1996||Uop||Process for removal of nitriles from etherification feedstocks|
|US7553995||Sep 11, 2007||Jun 30, 2009||Catalytic Distillation Technologies||Method of producing tertiary amyl ethyl ether|
|US7674941||Jun 13, 2008||Mar 9, 2010||Marathon Gtf Technology, Ltd.||Processes for converting gaseous alkanes to liquid hydrocarbons|
|US7838708||Jan 25, 2010||Nov 23, 2010||Grt, Inc.||Hydrocarbon conversion process improvements|
|US7847139||Jul 2, 2008||Dec 7, 2010||Grt, Inc.||Hydrocarbon synthesis|
|US7880041||Jul 16, 2007||Feb 1, 2011||Marathon Gtf Technology, Ltd.||Process for converting gaseous alkanes to liquid hydrocarbons|
|US7883568||Feb 5, 2007||Feb 8, 2011||Grt, Inc.||Separation of light gases from halogens|
|US7964764||Jan 7, 2010||Jun 21, 2011||Grt, Inc.||Hydrocarbon synthesis|
|US7998438||May 27, 2008||Aug 16, 2011||Grt, Inc.||Zone reactor incorporating reversible hydrogen halide capture and release|
|US8008535||Apr 30, 2008||Aug 30, 2011||Marathon Gtf Technology, Ltd.||Process for converting gaseous alkanes to olefins and liquid hydrocarbons|
|US8053616||Jul 1, 2009||Nov 8, 2011||Grt, Inc.||Continuous process for converting natural gas to liquid hydrocarbons|
|US8173851||Jun 3, 2009||May 8, 2012||Marathon Gtf Technology, Ltd.||Processes for converting gaseous alkanes to liquid hydrocarbons|
|US8198495||Mar 2, 2010||Jun 12, 2012||Marathon Gtf Technology, Ltd.||Processes and systems for the staged synthesis of alkyl bromides|
|US8232441||Jul 13, 2009||Jul 31, 2012||Marathon Gtf Technology, Ltd.||Process for converting gaseous alkanes to liquid hydrocarbons|
|US8273929||Jul 17, 2009||Sep 25, 2012||Grt, Inc.||Continuous process for converting natural gas to liquid hydrocarbons|
|US8282810||Jun 3, 2009||Oct 9, 2012||Marathon Gtf Technology, Ltd.||Bromine-based method and system for converting gaseous alkanes to liquid hydrocarbons using electrolysis for bromine recovery|
|US8367884||Feb 17, 2011||Feb 5, 2013||Marathon Gtf Technology, Ltd.||Processes and systems for the staged synthesis of alkyl bromides|
|US8415512||Oct 13, 2010||Apr 9, 2013||Grt, Inc.||Hydrocarbon conversion process improvements|
|US8415517||Jul 17, 2009||Apr 9, 2013||Grt, Inc.||Continuous process for converting natural gas to liquid hydrocarbons|
|US8436220||Jun 10, 2011||May 7, 2013||Marathon Gtf Technology, Ltd.||Processes and systems for demethanization of brominated hydrocarbons|
|US8642822||May 27, 2011||Feb 4, 2014||Marathon Gtf Technology, Ltd.||Processes for converting gaseous alkanes to liquid hydrocarbons using microchannel reactor|
|US8802908||Oct 8, 2012||Aug 12, 2014||Marathon Gtf Technology, Ltd.||Processes and systems for separate, parallel methane and higher alkanes' bromination|
|US8815050||Mar 22, 2011||Aug 26, 2014||Marathon Gtf Technology, Ltd.||Processes and systems for drying liquid bromine|
|US8829256||Jun 30, 2011||Sep 9, 2014||Gtc Technology Us, Llc||Processes and systems for fractionation of brominated hydrocarbons in the conversion of natural gas to liquid hydrocarbons|
|US8921625||Jul 17, 2009||Dec 30, 2014||Reaction35, LLC||Continuous process for converting natural gas to liquid hydrocarbons|
|US9133078||Dec 13, 2012||Sep 15, 2015||Gtc Technology Us, Llc||Processes and systems for the staged synthesis of alkyl bromides|
|US9193641||Dec 4, 2012||Nov 24, 2015||Gtc Technology Us, Llc||Processes and systems for conversion of alkyl bromides to higher molecular weight hydrocarbons in circulating catalyst reactor-regenerator systems|
|US9206093||Apr 17, 2014||Dec 8, 2015||Gtc Technology Us, Llc||Process for converting gaseous alkanes to liquid hydrocarbons|
|US20070251382 *||Feb 5, 2007||Nov 1, 2007||Gadewar Sagar B||Separation of light gases from halogens|
|US20080269534 *||Jul 2, 2008||Oct 30, 2008||Grt, Inc.||Hydrocarbon synthesis|
|US20080275284 *||May 20, 2008||Nov 6, 2008||Marathon Oil Company||Process for converting gaseous alkanes to liquid hydrocarbons|
|US20080314758 *||May 14, 2008||Dec 25, 2008||Grt, Inc.||Process for converting hydrocarbon feedstocks with electrolytic recovery of halogen|
|US20090005620 *||Jun 13, 2008||Jan 1, 2009||Marathon Gtf Technology, Ltd.||Processes for converting gaseous alkanes to liquid hydrocarbons|
|US20090069608 *||Sep 11, 2007||Mar 12, 2009||Boyer Christopher C||Method of producing tertiary amyl ethyl ether|
|US20100096588 *||Jul 17, 2009||Apr 22, 2010||Sagar Gadewar||Continuous Process for Converting Natural Gas to Liquid Hydrocarbons|
|US20100099929 *||Jul 17, 2009||Apr 22, 2010||Sagar Gadewar||Continuous Process for Converting Natural Gas to Liquid Hydrocarbons|
|US20100099930 *||Jul 17, 2009||Apr 22, 2010||Peter Stoimenov||Continuous Process for Converting Natural Gas to Liquid Hydrocarbons|
|US20100105972 *||Jan 7, 2010||Apr 29, 2010||Lorkovic Ivan M||Hydrocarbon Synthesis|
|US20100121119 *||Jan 25, 2010||May 13, 2010||Sherman Jeffrey H||Hydrocarbon Conversion Process Improvements|
|US20100270167 *||Apr 22, 2010||Oct 28, 2010||Mcfarland Eric||Process for converting hydrocarbon feedstocks with electrolytic and photoelectrocatalytic recovery of halogens|
|US20110034741 *||Oct 13, 2010||Feb 10, 2011||Sherman Jeffrey H||Hydrocarbon conversion process improvements|
|CN101245255B||Feb 14, 2007||May 30, 2012||中国石油化工股份有限公司||Etherification method for producing clean gasoline|
|U.S. Classification||208/67, 585/310, 585/312, 208/95, 208/176, 568/697, 208/78|
|International Classification||C10G11/18, C10G57/00|
|Cooperative Classification||C10G57/00, C10G11/18|
|European Classification||C10G57/00, C10G11/18|
|Mar 15, 1991||AS||Assignment|
Owner name: MOBIL OIL CORPORATION, A CORP. OF NY.
Free format text: ASSIGNMENT OF ASSIGNORS INTEREST.;ASSIGNOR:HARANDI, MOHSEN N.;REEL/FRAME:005677/0870
Effective date: 19910312
|Oct 1, 1996||REMI||Maintenance fee reminder mailed|
|Feb 23, 1997||LAPS||Lapse for failure to pay maintenance fees|
|May 6, 1997||FP||Expired due to failure to pay maintenance fee|
Effective date: 19970226